Catalytic decomposition of hydrocarbons



March 16, 1954 B. E. RoErHELl CATALYTIC VDECOMMF'OSITION OF HYDOCARBONSFiled March 15. 1949 )Bruno ERoetJfzeZL' :inventor v 5&5 ClbbrnegPatented Mar. 16, 1954 CATALYTIC DECOMPOSITION OF HYDR'OCARBON Sv BrunoE. Roetheli, Putney, England, `assignor to Standard Oil DevelopmentCompany, a corporation of Delaware Application March 15, 1949, SerialNo. 81,598

3 Claims.

The present invention rela-tes to the novel features hereinafter fullydisclosed in the specifi-cation and claims. It pertains to normallyendothermic selective decomposition reactions of hydrocarbons thatemploy a fluidized solid catalyst in an adiabatically heated reactionzone, for which such features of control are important.

This application is a continuation-impart of application, Serial Number555,456, led September 23, 1944, and nowabandoned.

The commercial selective vdecomposition process which involves aAdehydrogenation of hydrocarbons, e. g. dehydrogenation of parains toproduce desired olens or diolens such as butenes or butadiene, ordehydrogenation and rearrangement Without cracking are now a matter ofrecord and common practice. In the less modern commercial units builtfor dehydrogenation, the operation has been of the stationary bed type,that is to say, the catalyst is contained in a reaction vessel or casein the form of pills, pellets, extruded lengths, shaped bodies,granules, and

the like, and the mass of catalytic material is disposed in the form ofstationary beds. Of course, this type operation with a xed catalystgives difficulty in achieving vuniform reaction conditions, isnecessarily intermittent with respect to a given case or reactionvessel, for variations arise as to `points .of contact and catalystactivity and there comes a time when it is necessary to reactivate theentire catalyst to remove carbonaceous deposits contained thereon as aresult of the process.

More recently there has been work directed toward the end of producingolens by dehydrogenation in a continuous operation. The processesinvolved have used a powdered catalyst suspended in the reaction vaporsin thereaction zone in the form of a dense phase suspension. This typeof operation in which a delayed settler is employed has come to be knownas the fluid catalyst type of operation. With respect to the catalyticcracking of hydrocarbon oils, there has been put into commercialoperation the fluid catalyst type of operation, and in a preferredembodiment of this process a standpipe .in communication with anelevated regeneration vessel Vfeeds continuously by gravity and/ornatura-1 flow, a quantity of powdered catalyst to a reaction zonedisposed below the level of the regeneration zone in which reactionzone, by controlling the amount of' catalyst fed thereto and thevelocity of the vapors and/or gases flowing therethrough, thereiscaused' to be formedk the fluidized mass of dense suspensionpreviously referred to.

Although the use of `a fluidized :solid 'catalyst makes -a remarkableimprovement with regardlto uniformity of catalyst activity, elimination-of intermittent reactivation, and distribution of heat Within thereaction zone, there are -cert-a-in problems of temperature and :contacttime control that have become more accentuated in this type of processespecia-llywhen the heat Aof reaction substantial and when the reactionmust be closely controlled to obtain high selectivity towards 'formationof certain products.

In a preferred embodiment of the present invention, a system is used inwhich catalyst in powdered form is continuously discharged from anelevated regenerator through a standpipe into a reactor disposed at alower --point a hydrocarbon, to be selectively dehydrogenated, usually a-gasiform parafn or olen, being also forced into said reactor, the nowrate yof the gasif-orm paraffin 1or olefin being such that a densesuspension offcatalyst in the gaseous reactant is formed in thereactor.Catalyst of partly spent activity is wit-hdrawn from the reactorcontinuously and conveyed in a carrier gas stream to the regenerator.The catalyst `is regenerated with an oxidizing gas in the regeneratorand a stream of the regenerated catalyst is returned to the reactor thuscom-pleting a cycle of operations adapted for continuity.

The main object of the invention, is to supply heat chemically, or byexothermic reaction, in the reactor zone of a continuous dehydrogenationoperation of the type outlined above in such a way so as to haveexothermic heat substantially balance the endothermic heat absorbed bythe reaction of the hydrocarbon reactant.

Another important object is to control 'the adiabatic :heat input in thefluidized catalyst type of process, e. g. by having Ia propertemperature in the hydrocarbon feed in obtaining the desired heatbalance.

While 'it is known that iiuidized catalyst processes have Ibeen designedfor adding sensible heat to a reaction zone by heating the catalyst in azone outside the reaction zone, as for instance, by causing combustionof contaminants on fouled catalyst, burning added oil on the catalyst,letc., and thereafter transferring the uncooled catalyst to the reactionzone, there 4has apparently been no ragard hitherto for the propertemperature control of the hydrocarbon kfeed and proportioning of therecycled catalyst 'to the feed Afor 'attaining the objects set forthherein.

In fundamental experimental work on which this invention is based, itWas observed that -during a rluidized catalytic `del'nydrogenaton ofpa-raiiins or olefms, which are of course endothermic or heat absorbing,carbon oxides and Water vapor appear in the products issuing from thedehydrogenation zone. In such dehydrogenation operations, the catalyst,or at least an active component thereof, is usually a, partiallyreducible metallic oxide, such as molybdenum oxide, chromic oxide,tungstic oxide, or similar oxide of a heavy metal which can exist inmore than one state of oxidation and is difcult to reduce to theelementary metal While the partially reducible metallic oxides mightoxidize hydrogen, carbon, or hydrocarbons in the reaction zone toliberate heat, they also absorb heat in being reduced. At the same timeother possible causes of heat changes exist, such as heat of absorptionand desorption, hence it is diflicult to determine theoretically whatnet heat change occurs. However, by actual experimental determinationsthe net heat changes can be measured and the factors ascertained. Thusmeans are provided during the course of the dehydrogenation reaction forproviding in situ a quantity of chemical heat which shall be availableto compensate at least in part for the heat absorbed by the endothermicreaction therein taking place. The advantage of my process lies insubstituting chemical heat for Sensible heat which tends to minimize theprobability of local overheating when mixing hot catalyst from theregenerator with the feed stock and for obtaining an adiabatic operationwhich becomes more closely isothermal.

In the accompanying drawing is shown diagrammatically by means of asimplied flow plan the essential apparatus which illustrates a preferredembodiment of my invention.

Referring in detail to the drawing, I represents a reaction vessel orcase, and 3 represents a regeneration vessel, or case, provided with atorch oil inlet 4. These vessels are cylindrical shells having conicalor frusto-conical bases and crowns, and the regenerator is positionedabove the reactor so that, as will appear more fully hereinafter, anatural flow of powdered catalyst from the regenerator to the reactorthrough a standpipe may be effected. Assuming that the process to beoperated involves the dehydrogenation of normal butane, in operatingthis process group VI metal oxide supported on activated alumina may beused for instance, say, lO-io weight per cent of Cr2O3 carried on orsupported by 90-60 weight per. cent activated alumina. The catalyst isin the form of a powder having a particle size of from 100 to 400 mesh,but preferably 95 per cent of it is about 200 mesh.

Referring to the regenerator 3 which contains powdered catalystundergoing regeneration, a portion of this catalyst is continuouslywithdrawn through standpipe 5 controlled by a slide valve 6. Forpurposes of purging the catalyst of occluded oxygen and, at thesametime, inducing a free ilow without bridging or plugging, a fluidizinggas is bled into the standpipe 5 through a plurality of taps 'I such asflue gas, CO2, methane, or other inert gas, This iluidizing gas sweepsout entrapped cr occluded oxygen and also, as indicated, fluidizes thecatalyst so that it flows freely. At a point near the bottom of thestandpipe 5 the butane feed is injected through line I0. Additionaluidizing gas, steam, or the like may be injected through line Il. Thebutane gas mixes with the catalyst forming a suspension which isdischarged into the base of reaction vessel l and thence passes upwardlythrough a distributing grid G into the reaction zone proper. In thereaction vessel l the velocity of the gasiform material is controlledwithin the limits of from 1/2 to 5 ft. per second, preferably 1 to 3 ft.per second, to obtain the dense suspension previously referred to. Thedense suspension of uidized mass of catalyst in the gasiform material isobtained under the condition referred to by a sort of delayed settlingwherein the Velocity of the gas is sufficient to support the catalyst,but nevertheless permits a slippage so that the vapors pass through themass of catalyst and into a space in the crown of the reaction vesselwhere they are greatly depleted of catalyst. In other words, bycontrolling the amount of catalyst discharged into reaction vessel l andthe vapor gas velocity therein, upper level L of the dense phase of thesuspension is xed at a given point and the gases and/or vapors above thegiven point are greatly depleted of catalyst, or behold a smallproportion of the catalyst in dilute suspension. Thus, for example,where the density of the suspension in the reaction zone between G and Lmay be of the order of 15-35 lbs. per cu. ft. the gasiform materialactually issuing through line 20 from the reaction zone may have adensity as low as 0.03 lb. per cu. ft.

The usual dehydrogenation conditions are maintained in the reactionzone. Thus the temperature is of the order of S25-1050o F., the pressurearound 0-10 lbs. per square inch gauge, and the feed rate of hydrocarbonis such that there are from l@ to 2 pounds of hydrocarbon feed per hourper pound of catalyst in the reaction zone per hour, but preferablyaround 0.3-0.7 pound of hydrocarbon is fed per hour per pound ofcatalyst in the reactor. These operating conditions as to temperature,pressure and feed rate were known prior to this present invention, andno novelty is claimed in them per se, nor is novelty claimed in theseparation and recovery of the desired product, which in theillustration chosen is butylene. It will therefore be suicient to saythat the raw reaction products in line 20 are passed through one or moresolid-gas contacting devices, such as centrifugal or electricalseparators designed to remove catalyst fines from the raw product, whichseparated catalyst is then returned to the reaction zone via line 24,while the crude butylene passes into a purification system Where,according to the usual practice, it is separated from gases andunconverted parans by suitable fractionation and absorption. In somecases the overhead products can be used directly without purification.The advisability of this depends upon economics.

Referring again to the reaction zone, catalyst which becomes fouledduring the dehydrogenation or hydroforming operation is withdrawnthrough a standpipe 30 provided with a plurality of gas taps 32 intowhich purging gas such as flue gas may be injected into the standpipe 30for the lpurpose of displacing occluded volatile hydrocarbons and at thesame time increasing the fluidity of the downflowing fouled catalyst bypreventing bridging and plugging therein. The catalyst discharged into aline 40 containing an oxygen-containing gas, such as air, introducedinto line 39 to form a suspension which is pneumatically conveyed to theregenerator 3. In the regenerator the usual conditions of ternperatureand pressure are maintained to cause combustion of the carbonaceousdeposits in the fouled catalyst. Thus the temperature is controlledwithin the limits of, say. 1050 to 1200 F.

and under atmospheric or super-atmospheric pressure so as to effect thedesired purification of the catalyst by burning of the contaminants.Also, as inthe case of the reaction vessel l, the quantity of catalystfed to the regenerator and the velocity of the gasiform material in theregenerator is controlled so as to maintain within the regenerator by aprocess which amounts to delayed settling of a suspension, a densesuspension of catalyst in gas, having an upper level at L and beingconcentrated to the extent that between the grid G and the level L, themass of catalyst has a density of 15-40 lbs. per cubic foot depending onthe composition of the catalyst. Above L the concentration of catalystis greatly diminished so that the flue gases issuing through line 50have a density of the order of 0.03 lb. per cubic foot. To remove atleast a portion of the catalyst fines, the iiue gas is caused to owthrough a gas-solid contacting device 52y (or several of them) to effectseparation of catalyst from the flue gas, which catalyst is returned vialine 53 to the regeneration zone, while the flue gas is then passedthrough a heat recovery system 63 to recover at least a portion of itssensible heat. For simplicity sake, and also because many good methodsare known to the prior art, I have not shown in detail the conventionalwaste heat boilers,.heat exchangers, and the like, which may be employedto recover sensible heat from the hot flue gases, for the skilledengineer need not have this explained to him. Following regeneration,catalyst is returned through standpipe 5 to the reactor and the processis repeated. Unavoidably, of course, some catalyst may be lost from thesystem and therefore in order to maintain the catalyst inventory in thesystem at a constant value, and also to maintain the activity of thecatalyst at a desired level fresh catalyst in hopper 'I0 may be fedeither intermittently or continuously to the system through line 12.

In a general way a process illustrating a process to which amodification of the present invention can be applied now will bedescribed in detail to show how to accomplish the objectsaforementioned, namely, the adjustments of hydrocarbon feed temperaturesand catalyst to hydrocarbon feed ratios to obtain better temperature andheat control. In the first place, of course, the burning of thecontaminants on the catalyst in regeneration zone 3 provides as sensibleheat a. portion of the necessary heat for the main reaction taking placein reaction zone l, to supplement this quantity of heat commonly thehydrowith a rise'in vtemperature of the total system';

(catalyst-reactant). Metals of group VI form higher oxides in theregeneration step (oxidation) which in the presence of hydrocarbons arereduced to lower oxides with the formation of CO,

CO2 and H2O and a liberation of heat, i. e..thel

oxides furnish oxygen for formation of combustion products.

Intensive studies made of the reactions and rez oxides are difhcult toreduce to any lower state' of oxidation. Regeneration of such catalystwas found to be rapid `and complete, the carbon in the carbonaceous orcoke deposits being lowered by about 90% on contacting the spentcatalyst with oxygen for about 1 second. Analyses showedv that a portionof the oxygen oxidizes the catalyst..

From the results of heat balance measurementsf there was found to .be aheating effect by the regenerated catalyst in the reaction and thiseffect could not be accounted for simply by oxidation of hydrogen andcarbon.

Two different sets of heats of reaction were determined from a largeseries of runs. The first set was an evaluation of the overall heats ofreaction, which take into account the total heat input and heat loss, sothat these values include the heat evolved and absorbed by all reactionsoccurring. Such overall heats of reaction were found to vary from about770 to +790 (endothermic to exothermic) B. t. u. per pound of feedconverted. The second set of values were obtained by subtracting theheats of combustion of the carbon and hydrogen and should represent thenet heat of dehydrogenation of the feed.

Except for effects of variations in selectivity.. the net heats ofdehydrogenation should approximate the theoretical value, which is about900 B. t. u. per pound of butane converted. However, there were found tobe wide variations in the net heat of reaction values, ranging fromabout 1000 to +80 B. t. u. per pound of feed converted. In general, thetwo values, overall and net, for each run agree indicating that thefactors causing the variations are conditions of operation and that, thevariations were not due to haphazard inaccuracies. y

The data indicate that one factor having a certain effect is thecatalyst-to-hydrocarbon feed ratio. The effect of this factor is shownby the carbon feed has been heated to an elevated temlepesentive resultstabulated below.

Table I Catalyst to Butane Feed Ratio 6.1 6. 2 7. 2 Overall Heat ofReaction -510 -530 -440 Net Heat of Reaction -l, 010 -870 -900 Catalystto Butane Feed Ratio 8.0 9. 6 9.9 Overall Heat of Reaction 250 -130 NetHeat of Reaction.' 750 -620 655 *Interpolated Value.

perature within the reaction temperature range or close thereto forsupplying the endothermic reaction heat requirements.

vIn the course of operating a continuous dehydrogenation plant accordingto the previous description, catalyst containing partially reducibleoxides tend. to become oxidized in the regeneration step to a higherstate of oxidation and tend to become reduced to lower oxides in thereaction step. Reduction under certain conditions and The Table I datashows that when the catalyst to butane ratio is less than 8 parts byweight of catalyst to 1 part by weight of butane, the net heat ofdehydrogenation is practically constant. and approximates thetheoretical heat of dehydrogenation, 900 B. t. u. per 1b. of butaneconverted. When the ratio is increased to above 8:1, the net heat valuesare abruptly changed toward (exothermic) values which when plotted showa definite linear relationship to with certain reducing agents isaccomplished. reach apeak in the region where the ratio is amargo:

7 in :the range of. 16:1 to 18:1. At the' ratio of about 19:1. the netheat of reaction abruptly4 falls.

Thus it was indicated that. by hav'mg the catalyst to hydrocarbon ratiowithin the range of 8:1 to `21:1 the net heat of dehydrogenation is:substantially :counterbalanced by .exothermic heats of reaction, sothat less heat might be supplied to the butane feed .in preheating thefeed stream in a heated tube prior to its con tact with the `catalyst tolessen difficulties from thermal decompositiony .and avoid. excessivereaction when the hydrocarbon first contactsr the freshly reierctiyatedcatalyst, which is also at its highest temperature.

Actually vit was determined. from a series/of runs .that a .secondimportant factor in obtaining a smoother" operation, more erenv reactiontemperature control with substantial .heat input economy residesxin thetemperature of the preheated hydrocarbon feed as shown in representativedata tabulated-below:

[Climamio-alumina csatalysrl to butano rati-os .bctwcon Szl :nid 21:1.1

2 Temperatures F.'. l I

Regenerated Ca- I tal=yst............. Lilo 1,1138 l 1, 080 1,124 1,120

feed at inlet. 9S 27S 400 525 934 935 f l i y Hydrocarbon RcoctOrIUlet1'038 1,022 1,095 .1,030 .1, 072. 1, 080 Reactor Top. 1,005 .1,000 1,0411,005 992 900 Reactor vefraaa. '1,002 1,000 1,044 f 1,003 997 90S FeedConverted As illustrated Abythe data shown in Table 1I,

in general it was found that by keeping the temperature of thehydrocarbon feed low, the deviations in `temperature throughout thereaction zone is greatly diminished and a required reaction temperatureis readily maintained. From such data it was determined it isadvantagecus to avoid preheating the hydrocarbonfeed to above 700 'F'.

The following examples show the eiect of the catalystto hydrocarbon feedweight ratios atapproxi'mately minimum and maximum critical limits fromaverage runs on the heat evolved to compensate for the heat losses (tooutlet streams and from reaction apparatus by conduction and radiation)land for heat absorbed by the endothermic reaction.

.Table III Catalyst courpositon 10%(3220: on 90%.Allz0r Hydrocarbonlee-(l: Field butano 4lPositivevalucmeans overall boat of reaction 'isexothermic.

It can vbest 'be appreciatedfrom. the datain Table. III how by operatingwithin the .preferred limits vin which the catalyst Ito hydrocarbonratior is within the .range of 8 110.21, the reaction zone `can bemaintained 'at a relatively uniform temperature without preheating thefeedV to above.- 700" F., -thus presenting. thermal eracle-r ing in` anypreheatingof the feed. and preventing 8 l excessive reaction of thefresh feed with thel highly active freshly regenerated catalyst. Theaverage saving in thus preventing cracking` of the hydrocarbon reactantto lower hydrocarbons amounts vto about 9%.

By having the relatively low temperature hydrocarbon `feed streamcontact the freshly regenerated catalyst at a high temperature, beforethe Aresulting mixture is introduced into the reaction zone, the.hydrocarbon reactant stream precools the catalyst without beingimmediately subjected to .excessive reaction. The resulting reactionmixture then at a temperature Within the :optimum reaction temperatureenters .the reaction zone .and this is a contributing factor inpreventing an excessive temperature variation in the reaction zone.Forthe saine reason, it is desirable to .avoid having the regenerated:catalyst atan unduly high temperature, eyen though thisk preventsaddition of large amounts of heat on 4the feed .streams `to compensatefor heat losses and tosupply heat absorbed in the endothemnic reaction.By the controls of the present invention, the exothermic heat evolved inexcess of the heat needed in the endothermic-reaction can. in `part atleast compensate for these heat losses.

It is to be Iunderstood that vthe improvements of contacting vahydrocarbon reactant which isl to be catalytieally dehydrogenatedWithout thermal Acracking is applicable to paralns, monoolei'lns andnaphthenes having 4 to 16 carbon atoms per molecule, since suchhydrocarbons tend to undergo thermal cracking readily on being preheatedat temperatures close Yto their catalytic dehydrcgenation temperatures,which are generally in the range .of about 800 F.. to l00 F. Thus, .ingeneral these improvements are adapted for use in conducting endothermiccatalytic dehydrogenation and reforming re'- actions of. naphthahydrocarbons to form unsaturated and aromatic hydrocarbons.

@ther modifications of this invention are readily suggested to thoseskilled in the art wherein an oxidized catalyst in iluidized form froma. fre- :atcr `is contacted 'with a hydrocarbon reactant stream.

What is :claimed is:

l. In a continuous process of dehydrogenating a vaporiaed butano with ahot uidized catalyst that chrmniirrn oxide inits highestxstate of.oxidation on alumina, the steps which lcomprise mixing. a feed streamof 'the vaporized butano at a temperaturebetween 100 F. .and '100 F.with a streamer said hot; -fluidized catalyst at a temperature below1200 F. but above lthe .re-v

action temperature range of 1000-1050 F. to form a reaction mixture ofbetween 8 and 12 parts by weight of said catalyst suspended in one partby we'ghtof the vaporired butano, substantially all .the catalyst thusadmixed having been freshly regenerated by oxidation at 1050 F. to 1200F. continuously passing said reaction mixture into `a reactionzonewherein the catalyst forms a dense nuidized bed, continuouslywithdrawing gaseous reaction products 4andxspent catalystzf rom saidreactim zone While: thusv keeping the ter peratures throughout thereaction zone from vvrarying by more than about 50 Fahrenheit degrecs.

2'. In a continuous process-of dehydrogenating a. yaporiaed aCi. Yto C15'hydrocarbon reactant. .of the .group consisting `of paraffins-xmono-olens, and naphthenes-with a hot iuidized catalyst ofchromieoxidefonalumina. the steps' which cornprise mixing a `feed-stream .of the hydrocarbon heat of the dehydrogenation reaction by asuiiicient quantity of exothermic heat liberated in the reaction zone tothereby maintain a more uniform temperature varying by less than 50 F.Within the limits of 925 F. and 1050 F. throughout the reaction zone,withdrawing from the reaction zone a stream of the resultingdehydrogenated hydrocarbon product, and continuously passing thecatalyst used in the reaction zone to a regeneration zone for oxidationat 1050 to 1200 F.

3. In a continuous process of dehydrogenating a vaporized C4 to C16paraffin hydrocarbon reactant with a hot uidized catalyst of chromicoxide on alumina, the steps which comprise admixing said catalyst at a.temperature in the range of 1050 F. to 1200 F. with a stream of thevaporized hydrocarbon reactant having a temperature in the range of 100F. to 700 F. to form a continuous reaction mixture stream having a 10temperature in the range of 1000 F. to 1050 F., substantially all ofsaid adrnixed catalyst having been freshly regenerated by oxidation at1050 F. to 1200 F., passing said reaction mixture stream into a reactionzone which loses heat carried away by gaseous products and spentcatalyst continuously withdrawn therefrom, by conduction, radiation andendothermic dehydrogenation of the parain hydrocarbon reactant,generating excthermic heat within the reaction zone at substantially thesame rate which the reaction zone loses heat to keep temperaturesthroughout the reaction zone within the range of 1000 F. to 1050 F. byhaving substantially from 9.6 to 19 parts by Weight of the hot catalystmixed continuously with each one part by Weight of the hydrocarbonreactant entering said reaction zone, withdrawing gaseous products fromthe reaction zone, and continuously passing the catalyst from thereaction zone to a regeneration zone Where the catalyst is reoxidized at1050 F. to 1200 F.

BRUNO E. ROETI-IELI.

References Cited in the le of this patent UNITED STATES PATENTS NumberName Date 2,231,231 Subkow Feb. 11, 1941 2,327,175 Conn Aug. 17, 19432,392,248 Layng et a1. Jan. 1, 1946 2,397,352 Hemminger Mar. 26, 19462,403,375 Kassel July 2, 1946

1. IN A CONTINOUS PROCESS OF DEHYDROGENATING A VAPORIZED BUTANE WITH AHOT FLUIDIZED CATALYST THAT CONTAINS CHROMIUM OXIDE IN ITS HIGHEST STATEOF OXIDATION ON ALUMINA, THE STEPS WHICH COMPRISE MIXING A FEED STREAMOF THE VAPORIZED BUTANE AT A TEMPERATURE BETWEEN 10/* F. AND 700* F.WITH A STREAM OF SAID HOT FLUIDIZED CATALYST AT A TEMPERATURE BELOW1200* F. BUT ABOVE THE REACTION TEMPERATURE RANGE OF 1000-1050* F. TOFORM A REACTION MIXTURE OF BETWEEN 8 AND 12 PARTS BY WEIGHT OF SAIDCATALYST SUSPENDED IN ONE PART BY WEIGHT OF THE VAPORIZED BUTANE,SUBSTANTIALLY ALL THE CATALYST THUS ADMIXED HAVING BEEN FRESHLYREGENERATED BY OXIDATION AT 1050* F. TO 1200* F. CONTINUOUSLY PASSINGSAID REACTION MIXTURE INTO A REACTION ZONE WHEREIN THE CATALYST FORMS ADENSE FLUIDIZED BED, CONTINUOUSLY WITHDRAWING GASEOUS REACTION PRODUCTSAND SPENT CATALYST FROM SAID REACTION ZONE WHILE THUS KEEPING THETEMPERATURES THROUGHOUT THE REACTION ZONE FROM VARYING BY MORE THANABOUT 50 FAHRENHEIT DEGREES.